Isomerization of Benzene-Containing Feedstocks

ABSTRACT

The benzene content in a gasoline pool is reduced by a process that hydrogenates a benzene-containing isomerization zone feedstream. The additional cyclic hydrocarbons produced by the saturation of benzene can be processed in the isomerization zone for ring opening to increase the available paraffinic feedstock or the isomerization zone can be operated to pass the cyclic hydrocarbons through to a product recovery section. The isomerization zone feedstream is treated to remove contaminants and dried before entering the hydrogenation zone.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Continuation of copending application Ser. No.11/750,521 filed May 18, 2007, the contents of which are herebyincorporated by reference in its entirety.

FIELD OF THE INVENTION

This invention relates generally to the isomerization of hydrocarbons.This invention relates more specifically to the processing ofbenzene-containing hydrocarbon feeds and the isomerization of lightparaffins.

BACKGROUND OF THE INVENTION

High octane gasoline is required for modern gasoline engines. Benzenehas a high octane number value and has been previously blended intogasoline. However, as benzene is phased out of gasoline forenvironmental reasons, it has become increasingly necessary to rearrangethe structure of the hydrocarbons used in gasoline blending in order toachieve high octane ratings. Catalytic reforming and catalyticisomerization are two widely used processes for this upgrading.

A gasoline blending pool is usually derived from naphtha feedstocks andincludes C₄ and heavier hydrocarbons having boiling points of less than205° C. (395° F.) at atmospheric pressure. This range of hydrocarbonincludes C₄-C₉ paraffins, cycloparaffins and aromatics. Of particularinterest have been the C₅ and C₆ normal paraffins which have relativelylow octane numbers. The C₄-C₆ hydrocarbons have the greatestsusceptibility of octane improvement by lead addition and were formerlyupgraded in this manner. Octane improvement can also be obtained bycatalytically isomerizing the paraffinic hydrocarbons to rearrange thestructure of the paraffinic hydrocarbons into branch-chained paraffinsor reforming to convert the C₆ and heavier hydrocarbons to aromaticcompounds. Normal C₅ hydrocarbons are not readily converted intoaromatics, therefore, the common practice has been to isomerize theselighter hydrocarbons into corresponding branch-chained isoparaffins.Although the non-cyclic C₆ and heavier hydrocarbons can be upgraded intoaromatics through dehydrocyclization, the conversion of C₆'s toaromatics creates higher density species and increases gas yields withboth effects leading to a reduction in liquid volume yields. Therefore,it is preferable to charge the non-cyclic C₆ paraffins to anisomerization unit to obtain C₆ isoparaffin hydrocarbons. Consequently,octane upgrading commonly uses isomerization to convert normal C₆ andlighter boiling hydrocarbons and reforming to convert C₆ cycloparaffinsand higher boiling hydrocarbons.

In the reforming processing, C₆ cycloparaffins and other higher boilingcyclic hydrocarbons are converted to benzene and benzene derivatives.Since benzene and these derivatives have a relatively high octane value,the aromatization of these naphthenic hydrocarbons has been thepreferred processing route. However, many countries are contemplating orhave enacted legislation to restrict the benzene concentration of motorfuels. Therefore, processes are needed for reducing the benzene contentof the gasoline pool while maintaining sufficient conversion to satisfythe octane requirements of modern engines.

Combination processes using isomerization and reforming to convertnaphtha range feedstocks are well known. U.S. Pat. No. 4,457,832 usesreforming and isomerization in combination to upgrade a naphthafeedstock by first reforming the feedstock, separating a C₅-C₆ paraffinfraction from the reformate product, isomerizing the C₅-C₆ fraction toupgrade the octane number of these components and recovering a C₅-C₆isomerate liquid which may be blended with the reformate product. U.S.Pat. No. 4,181,599 and U.S. Pat. No. 3,761,392 show a combinationisomerization-reforming process where a full range naphtha boilingfeedstock enters a first distillation zone which splits the feedstockinto a lighter fraction that enters an isomerization zone and a heavierfraction that is charged as feed to a reforming zone. In both the '392and '599 patents, reformate from one or more reforming zones undergoesadditional separation and conversion, the separation including possiblearomatics recovery, which results in additional C₅-C₆ hydrocarbons beingcharged to the isomerization zone.

The benzene contribution from the reformate portion of the gasoline poolcan be decreased or eliminated by altering the operation of thereforming section. There are a variety of ways in which the operation ofthe reforming section may be altered to reduce the reformate benzeneconcentration. Changing the cut point of the naphtha feed split betweenthe reforming and isomerization zones from 82 to 93° C. (180° to 200°F.) will remove benzene, cyclohexane and methylcyclopentane from thereformer feed. Benzene can alternately also be removed from thereformate product by splitting the reformate into a heavy fraction and alight fraction that contains the majority of the benzene. Practicingeither method will put a large quantity of benzene into the feed to theisomerization zone.

The isomerization of paraffins is a reversible reaction which is limitedby thermodynamic equilibrium. The basic types of catalyst systems thatare used in effecting the reaction are a hydrochloric acid promotedaluminum chloride system and a supported aluminum chloride catalyst.Either catalyst is very reactive and can generate undesirable sidereactions such as disproporationation and cracking. These side reactionsnot only decrease the product yield but can form olefinic fragments thatcombine with the catalyst and shorten its life. One commonly practicedmethod of controlling these undesired reactions has been to carry outthe reaction in the presence of hydrogen. With the hydrogen that isnormally present and the high reactivity of the catalyst, any benzeneentering the isomerization zone is quickly hydrogenated. Thehydrogenation of benzene in the isomerization zone increases theconcentration of napthenic hydrocarbons in the isomerization zone.

It has been discovered that placing a hydrogenation reaction zone infront of an isomerization reaction zone but downstream of the feeddriers required for the isomerization catalyst allows savings byreduction of equipment count and cost as well as a reduction in theamount of hydrogen required for the process. Placing the hydrogenationreaction zone downstream of the feed driers, allows the productcondensers and receiver that would normally be required downstream ofthe hydrogenation reactor to be eliminated. Because the receiver hasbeen eliminated, there is no hydrogen venting required. Hydrogen is avaluable commodity to refiners who are in need of ways to reducehydrogen usage. Furthermore, in the present invention, low pressure feeddriers may be used. Driers that are only operated at low pressures areless costly than high pressure driers and the cost of the many valvesassociated with the driers for the purposes of regenerating the driersieves is reduced significantly for low pressure driers. Finally,additional utility savings are realized by the elimination of thecondensing equipment normally required downstream of the hydrogenationreaction zone.

SUMMARY OF THE INVENTION

This invention is a process for converting a feedstock comprising C₄-C₇paraffins and C₅-C₇ cyclic hydrocarbons including benzene. Thisinvention uses a hydrogenation zone upstream of the isomerizationreactors to saturate benzene and simultaneously heat the feed to theisomerization zone. The use of a separate hydrogenation zone also lowersthe overall temperature of the isomerization zone feed as the benzene issaturated—lower temperatures minimize undesirable hydrocrackingreactions. Also performing the highly exothermic benzene saturationreaction in a lead reactor that has a lower temperature reduces thecoking that could occur in the isomerization zone as a result of thehigher overall temperatures.

Accordingly in one embodiment, this invention is a process for theisomerization of a C₄-C₆ paraffinic feedstock that contains at least 1wt.-% benzene. The process includes the steps of combining the feedstockwith a hydrogen-rich gas stream to produce a combined feed. The combinedfeed is passed to a hydrogenation zone and contacted therein with ahydrogenation catalyst to saturate benzene and heat the feedstream. Thesaturated feedstream is recovered from the hydrogenation zone and has abenzene concentration of less than 1.5 wt.-%. At least a portion of thesaturated feedstream is passed from the hydrogenation zone to anisomerization zone without heating and contacted with an isomerizationcatalyst at isomerization conditions.

In a yet further embodiment, this invention is a process for theisomerization of C₅-C₆ paraffinic feedstock that contain at least 1wt.-% benzene. The process dries the feedstock before combining thefeedstock with a dried hydrogen-rich gas to produce a combined feed thatis passed at a temperature of from 38 to 232° C. (100 to 450° F.) to anhydrogenation zone and contacted therein with a hydrogenation catalyst.In another embodiment the temperature of the combined feed is 127 to232° C. (260 to 450° F.) or 149 to 204° C. (300 to 400° F.). To heat thecombined feed, the combined feed may be heat exchanged withisomerization and hydrogenation zone effluents. Contact with thehydrogenation catalyst saturates the benzene and the exothermic reactionheats the saturated feedstream (hydrogenation zone effluent) to atemperature of from 149 to 288° C. (200 to 450° F.). In anotherembodiment the saturated feedstream is heated to 177 to 274° C. (350 to525° F.) or 204 to 274° C. (400 to 525° F.). The saturated feedstreamhas a benzene concentration of from about 0.01 to about 5 wt-% or fromabout 0.1 to about 1.5 wt.-% and is heat exchanged with the combinedfeed and the feedstock, and possibly cooled, before being passed to anisomerization zone. The saturated feedstream is contacted with anisomerization catalyst in the isomerization zone to isomerize C₅-C₆hydrocarbons. An isomerate product essentially free of benzene isrecovered from the isomerization zone. Downstream separations may beused to recycle low octane components of the isomerization zoneeffluent.

Other embodiments, aspects and details of this invention are disclosedin the following detailed description of the invention.

BRIEF DESCRIPTION OF THE DRAWING

The FIGURE shows a schematic flow diagram of one embodiment of theprocess.

DETAILED DESCRIPTION OF THE INVENTION

A basic arrangement for the processing equipment used in this inventioncan be readily understood by a review of the flow scheme presented inthe FIGURE. The FIGURE does not show all pumps, condensers, reboilers,instruments and other well-known items of processing equipment in orderto simplify the drawing. However, the discussion points out severalitems of traditional processing equipment that may be eliminated andthus provide both a capital cost savings and an operational costsavings.

Looking at the FIGURE, a feedstream comprising at least C₅ and C₆paraffins along with at least 1 wt.-% benzene enter the process throughline 10 and pass through a sulfur guard bed 12 that removes sulfur fromthe feedstream. The sulfur-depleted feedstream in line 13 is passedthrough a low pressure drier 11 to remove water. It is important to notethat line 13 is not passed through a reactor, nor is hydrogen added,before being dried in low pressure drier 11. This eliminates the needfor product condensers and a receiver on line 13. The elimination of acommonly used condenser provides an operational and equipment costsavings, and the elimination of the receiver additionally eliminates theneed for a hydrogen vent. Hydrogen is a valuable component in refineriestoday, and conservation of hydrogen results in positive value for therefiner. Finally, only a low pressure drier 11 is required. Highpressure driers and their associated valves are far more costly than lowpressure driers and their associated valves. Thus a cost savings isrealized in requiring only a low pressure drier as opposed to a highpressure drier.

Make-up hydrogen enters the process through line 14 and passes through adrier 16 for removal of water and sulfur. The dried feedstream in line15 and the dried hydrogen from line 17 are combined in line 18 to form acombined feed. The combined feed 18 is heat exchanged in an exchanger 24against the contents of line 20 which carries the effluent from a secondisomerization reactor 22. The contents of line 18 are further heatexchanged in a heat exchanger 26 against the contents of line 28 whichcarries the effluent from a first isomerization reactor 30. The contentsof line 18 are still further heat exchanged in a heat exchanger 25against the contents of line 34 which carries the effluent from adehydrogenation reactor 32. The hydrogenation reactor 32 receives thecontents of line 18, the combined feed. The hydrogenation reactorsaturates benzene present in the combined feed and further heats thecombined feed. Line 34 carries a saturated feed from hydrogenationreactor 32 to the first isomerization reactor 30. A chloride-containingcompound is injected into the contents of line 34 by a line 80.

A first stage of isomerization takes place in reactor 30. Following thefirst stage of isomerization, the effluent in line 28 is exchanged inheat exchanger 26 against the combined feed in line 18 as discussedabove. Line 28 then carries the partially cooled isomerization effluentfrom reactor 30 to reactor 22. After further isomerization in reactor22, an isomerate product is taken by line 20, heat exchanged against thecombined feed in line 18 using heat exchanger 24 and then is passed to afractionation column 38. Fractionation column 38 removes light gasesfrom the isomerate product which are taken overhead by line 42 andwithdrawn from the process through the top of a receiver 44 via line 50.Recycle is conducted back to fractionation column 38 via line 46. Thestabilized isomerate product is withdrawn from the bottom offractionation column 38 by line 40. To increase efficiency, stabilizedisomerate product in line 40 is conducted to deisohexanizer 58 toseparate low octane alkanes, such as normal or single branchedisoparaffins and cyclic compounds such as cyclohexane, for recycle tothe isomerization zone via line 64. Valuable isomerate product in lines60 and 62 are combined into final product 66.

Suitable feedstocks for this invention will include C₄ plus hydrocarbonsup to an end boiling point of about 250° C. (482° F.). The feedstocksthat are used in this invention will typically include hydrocarbonfractions rich in C₄-C₆ normal paraffins. The term “rich” is defined tomean a stream having more than 50% of the mentioned component. Inaddition, the feedstock will include significant amounts of benzene. Theconcentration of benzene in the feedstock will at least equal 1.0 wt.-%and will normally be higher. The concentration of benzene may be from 1to 25 wt.-%, and is expected to usually be in the range of 3 to 15 wt.-%or 5 to 12 wt.-%. The other feed components will usually comprise C₅-C₆cyclic and paraffinic hydrocarbons with normal pentane, normal hexane,and isohexane providing most of the paraffinic components. Wheremultiple streams are combined to form a feedstock, the benzene in one ofthe feeds may be much higher than 25 wt.-%. The dilution effect ofcombining the streams results in the benzene being at a manageablelevel.

The isomerization zone and hydrogenation zone catalysts are often sulfursensitive. Suitable guard beds or adsorptive separation processes may beused to reduce the sulfur concentration of the feedstock. The FIGUREshows the treatment of the feedstock to remove sulfur upstream of thefeedstock drier, hydrogen addition point, and the hydrogenation zone. Itis important that the sulfur guard bed be located upstream of the driersince water may be liberated from fresh guard bed adsorbent. The feedstream is heated by heat exchange with the effluent of the benzenesaturation reactor using a heat exchanger before being passed to thesulfur guard bed. If needed, additional heat may be input into stream 10before reaching sulfur guard bed 12. The feed stream may be heated withany suitable process stream such as the stabilizer bottoms or with autility stream such as steam or hot oil.

Also, some of the possible isomerization zone catalysts suitable for usein this invention are highly sensitive to water and other contaminants.In order to keep water content within acceptable levels for suchcatalysts, the streams directed to the isomerization zone are firstpassed through at least one drier. The drier for this purpose reduceswater content to 0.1 ppm or less, and suitable adsorption processes forthis purpose are well known in the art. The specific placement of thedriers in relation to the guard beds and other streams allows for lowpressure driers to be used to dry the feedstream. Low pressure driersand their associated regeneration switching valves are much less costlythan high pressure driers and are less costly to operate as well. Asshown in the FIGURE, the feedstock passes through a drier and thehydrogen stream passes through another drier before the feedstock andthe hydrogen stream are combined to form the combined feed. It isimportant to note that both the feedstock drier and the hydrogen drierare upstream of the hydrogenation zone.

This specific arrangement results in the elimination of the need foradditional equipment such as a condenser, a receiver and a high pressuredrier on the feedstock stream, and makes the most use of equipmentcommonly found in existing systems built at a time where there was lessof a need to process benzene. Previous isomerization processes, wherethe benzene concentration in the feedstock was less than 5 wt.-%, oftenhad a drier that operated at low pressure. In contrast, previousisomerization processes with higher concentrations of benzene in thefeedstock, first saturated the benzene and then dried the product beforepassing the stream to the isomerization zone. Since the hydrogenatedproduct stream was two-phase, a condenser and a receiver were requiredto provide a liquid stream that was sent to the driers. To avoid theloss of isopentanes in the receiver off-gas, the receiver and thereforethe subsequent drier was operated at high pressure and a more costlyhigh pressure drier was required. With the current need for existingprocesses to be revised to process feedstocks containing higher benzeneconcentrations, and the existing low pressure driers, a novel flowscheme was required to allow the processing of a benzene containingfeedstock while at the same time only requiring a low pressure drier. Inaddition to the capital and operating cost savings associated with theelimination of the condenser and receiver and the ability to reuseexisting low pressure feed driers, this flow scheme eliminates the needfor sulfur guard beds on the hydrogen stream sent to the saturationreactor. All of the hydrogen used for the hydrogenation andisomerization zones is sent through hydrogen driers where both sulfurand water contained in the hydrogen stream are removed.

A hydrogen stream is combined with the feedstock to provide hydrogen forthe hydrogenation and isomerization zones. When the hydrogen is addeddownstream of the feedstock treating section, the hydrogen stream alsoundergoes drying or other treatment, such as sulfur removal, necessaryfor the sustained operation of the isomerization zone or hydrogenationzone. The hydrogenation of benzene in the hydrogenation zone results ina net consumption of hydrogen. Although hydrogen is not consumed by theisomerization reaction, the isomerization of the light paraffins isusually carried out in the presence of hydrogen. Therefore, the amountof hydrogen added to the feedstock should be sufficient for both therequirements of the hydrogenation zone and the isomerization zone.

The amount of hydrogen admixed with the feedstock varies widely. For theisomerization zone alone, the amount of hydrogen can vary to produceanywhere from a 0.01 to a 10 hydrogen to hydrocarbon ratio in theisomerization zone effluent. Consumption of hydrogen in thehydrogenation zone increases the required amount of hydrogen admixedwith the feedstock. The input through the hydrogenation zone usuallyrequires a relatively high hydrogen to hydrocarbon ratio to provide thehydrogen that is consumed in the saturation reaction. Therefore,hydrogen will usually be mixed with the feedstock in an amountsufficient to create a combined feed having a hydrogen to hydrocarbonratio of from 0.1 to 2. Lower hydrogen to hydrocarbon ratios in thecombined feed are preferred to simplify the system and equipmentassociated with the addition of hydrogen. At minimum, the hydrogen tohydrocarbon ratio must supply the stoichiometric requirements for thehydrogenation zone. In order for the hydrogenation zone to operate atthe mild conditions of this invention, it is preferable that an excessof hydrogen be provided with the combined feed. Although no net hydrogenis consumed in the isomerization reaction, the isomerization zone willhave a net consumption of hydrogen often referred to as thestoichiometric hydrogen requirement which is associated with a number ofside reactions that occur. These side reactions include saturation ofolefins and aromatics, cracking and disproportionation. Due to thepresence of the hydrogenation zone, little saturation of olefins andaromatics will occur in the isomerization zone. Nevertheless, hydrogenin excess of the stoichiometric amounts for the side reactions ismaintained in the isomerization zone to provide good stability andconversion by compensating for variations in feedstream compositionsthat alter the stoichiometric hydrogen requirements and to prolongcatalyst life by suppressing side reactions such as cracking anddisproportionation. Side reactions left unchecked reduce conversion andlead to the formation of carbonaceous compounds, i.e., coke, that foulthe catalyst. As a result, the effluent from the hydrogenation zoneshould contain enough hydrogen to satisfy the hydrogen requirements forthe isomerization zone. In one embodiment the effluent from thehydrogenation zone has a hydrogen to hydrocarbon mole ratio of fromabout 0.05 to about 2, in another embodiment the ratio is about 0.1 toabout 1.5 and in yet another embodiment the ratio is about 0.1 to 1.0.

It has been found to be advantageous to minimize the amount of hydrogenadded to the feedstock. When the hydrogen to hydrocarbon ratio at theeffluent of the isomerization zone exceeds about 0.20, it is noteconomically desirable to operate the isomerization process without therecovery and recycle of hydrogen to supply a portion of the hydrogenrequirements. Facilities for the recovery of hydrogen from the effluentare needed to prevent the loss of product and feed components that canescape with the flashing of hydrogen from the isomerization zoneeffluent. These facilities add to the cost of the process and complicatethe operation of the process. The isomerization zone can be operatedwith the effluent hydrogen to hydrocarbon ratio as low as 0.05 withoutadversely affecting conversion or catalyst stability. Accordingly wherepossible, the addition of hydrogen to the feedstock will be kept tobelow an amount that will produce a hydrogen to hydrocarbon ratio inexcess of 0.20 in the effluent from the isomerization zone.

The combined feed in line 18 comprising hydrogen and the feedstock enterthe hydrogenation zone. The hydrogenation zone is designed to saturatebenzene at relatively mild conditions. The hydrogenation zone comprisesa bed of catalyst for promoting the hydrogenation of benzene. Examplesof catalyst compositions include platinum group, tin or cobalt andmolybdenum metals on suitable refractory inorganic oxide supports suchas alumina. In one embodiment, the alumina is an anhydrous gamma-aluminawith a high degree of purity. The term platinum group metals refers tonoble metals excluding silver and gold which are selected from the groupconsisting of platinum, palladium, germanium, ruthenium, rhodium,osmium, and iridium.

Such catalysts have been found to provide satisfactory benzenesaturation at conditions including temperatures as low as 38° C. (100°F.), pressures from 1400 to 4800 kPa(g) (200 to 700 psig), an inlethydrogen to hydrocarbon ratio in the range of 0.1 to 2, and a 1 to 40liquid hourly space velocity (LHSV). Other suitable pressures includefrom about 2068 to about 4137 kPa(g) (300 to 600 psig) and from about2413 to about 3792 kPa(g) (350 to 550 psig) and other suitable liquidhourly space velocities include from about 4 to about 20 and about 8 toabout 20 hr⁻¹. In another embodiment of this invention, the feedentering the hydrogenation zone will be heated to a temperature in therange of 38 to 232° C. (100 to 450° F.), 127 to 232° C. (260 to 450° F.)or 149 to 204° C. (300 to 400° F.) by heat exchange with the effluentfrom the hydrogenation and isomerization zones. The exothermicsaturation reaction increases the heat of the combined feed andsaturates essentially all of the benzene contained therein. The effluentfrom the hydrogenation zone provides a saturated feed for theisomerization zone that will typically contain from 0.01 wt.-% to 5wt.-% or from 0.1 wt.-% to 1.5 wt.-% benzene or from 0.1 to 1.0 wt.-%benzene.

With the hydrogenation reaction being exothermic, the saturated feedfrom the hydrogenation reactor is typically at a temperature in therange of 149 to 288° C. (200 to 550° F.); 177 to 274° C. (350 to 525°F.); or 204 to 274° C. (400 to 525° F.). The isomerization zone operatesat a lower temperature range, so the heat of the saturated feed may berecovered and used to provide heat to other colder streams either withinthe process or from outside the process. For example, the saturated feedmay be heat exchanged with the combined feed and with the feedstock. Ifthe saturated feed is still too high in temperature even after heatexchange, the saturated feed may be cooled using conventionaltechniques.

Saturated feed from the hydrogenation zone enters the isomerization zonefor the rearrangement of the paraffins contained therein from lesshighly branched hydrocarbons to more highly branched hydrocarbons.Furthermore, if there are any unsaturated compounds that enter theisomerization zone after passage through the hydrogenation zone, theseresidual amounts of unsaturated hydrocarbons will be quickly saturatedin the isomerization zone. The isomerization zone uses a solidisomerization catalyst to promote the isomerization reaction. There area number of different isomerization catalysts that can be used for thispurpose. The zeolitic type isomerization catalysts are well known andare described in detail in U.S. Pat. No. 3,442,794 and U.S. Pat. No.3,836,597. Other catalysts include those such as described in U.S. Pat.No. 6,927,188.

The high chloride catalyst on an alumina base that contains platinum isalso well known in the art and not described in detail here. This typeof catalyst also contains a chloride component. The chloride componenttermed in the art “a combined chloride” is present in an amount fromabout 2 to about 10 wt.-% based upon the dry support material.

It is generally known that high chlorided platinum-alumina catalysts ofthis type are highly sensitive to sulfur and oxygen-containingcompounds. Therefore, the feedstock must be relatively free of suchcompounds. A sulfur concentration no greater than 0.1 ppm is generallyrequired at the reactor inlet. The presence of sulfur in the feedstockserves to temporarily deactivate the catalyst by platinum poisoning.Activity of the catalyst may be restored by hot hydrogen stripping ofsulfur from the catalyst composite or by lowering the sulfurconcentration in the reactor feed to below 0.1 ppm so that thehydrocarbon will desorb the sulfur that has been absorbed on thecatalyst. Water can act to permanently deactivate the catalyst byremoving high chloride from the catalyst and replacing it with inactivealuminum hydroxide. Therefore, water, as well as oxygenates, inparticular C₁-C₅ oxygenates, that can decompose to form water, can onlybe tolerated in very low concentrations. In general, this requires alimitation of oxygenates in the feed to about 0.1 ppm or less. Aspreviously mentioned, the feedstock may be treated by any method thatwill remove water and sulfur compounds. Sulfur may be removed from thefeedstock by hydrotreating. Adsorption processes for the removal ofsulfur and water from hydrocarbon streams are also well known to thoseskilled in the art.

Operating conditions within the isomerization zone are selected tomaximize the production of isoalkane product from the feed components.Inlet temperatures to, and temperatures within the reaction zone willusually range from about 38° C. to about 260° C. (100° F. to 500° F.) or104° C. to 204° C. (220° F. to 400° F.) or 104° C. to 177° C. (220° F.to 350° F.). Lower reaction temperatures are preferred for purposes ofisomerization conversion since they favor isoalkanes over normal alkanesin equilibrium mixtures. The isoalkane product recovery can be increasedby opening some of the cyclohexane rings produced by the saturation ofthe benzene. However, if it is desired, maximizing ring opening usuallyrequires temperatures in excess of those that are most favorable from anequilibrium standpoint. For example, when the feed mixture is primarilyC₅ and C₆ alkanes, temperatures in the range of 60° to 160° C. aredesired from a normal-isoalkane equilibrium standpoint but, in order toachieve significant opening of C₅ and C₆ cyclic hydrocarbon ring, thepreferred temperature range for this invention lies between 100° to 200°C. When it is desired to also isomerize significant amounts of C₄hydrocarbons, higher reaction temperatures are required to maintaincatalyst activity. Thus, when the feed mixture contains significantportions of C₄-C₆ alkanes the most suitable operating temperatures forring opening and isoalkane equilibrium coincide and are in the rangefrom 145° to 225° C. The reaction zone may be maintained over a widerange of pressures. Pressure conditions in the isomerization of C₄-C₆paraffins range from 1380 kPa(g) to 4830 kPa(g) (200 to 700 psig).Higher pressures favor ring opening, therefore, embodiments may usepressures for this process in the range of from 2410 kPa(g) to 4830kPa(g) (350 to 700 psig) when ring opening is desired. The feed rate tothe reaction zone can also vary over a wide range. These conditionsinclude liquid hourly space velocities ranging from 0.5 to 12 hr⁻¹, orbetween 0.5 and 3 hr⁻¹.

Depending upon the catalyst selected, operation of the reaction zone mayalso require the presence of a small amount of an organic chloridepromoter. The organic chloride promoter serves to maintain a high levelof active chloride on the catalyst as small amounts of chloride arecontinuously stripped off the catalyst by the hydrocarbon feed. Theconcentration of promoter in the reaction zone is usually maintained atfrom 30 to 300 ppm. Suitable promoter compounds include oxygen-freedecomposable organic chlorides such as perchloroethylene, carbontetrachloride, proplydichloride, butylchloride, and chloroform to nameonly a few of such compounds. The addition of chloride promoter afterthe hydrogenation reactor, as shown in the FIGURE, may be carried out atsuch a location to expose the promoter to the highest availabletemperature and assure its complete decomposition. The need to keep thereactants dry is reinforced by the presence of the organic chloridecompound which may convert, in part, to hydrogen chloride. As long asthe process streams are kept dry, there will be no adverse effect fromthe presence of small amounts of hydrogen chloride.

A preferred manner of operating the process is in a two-reactor,reaction zone system. The catalyst used in the process can bedistributed equally or in varying proportions between the two reactors.The use of two reaction zones permits a variation in the operatingconditions between the two reaction zones to enhance isoalkaneproduction. The two reaction zones can also be used to perform cyclichydrocarbon conversion in one reaction zone and normal paraffinisomerization in the other. In this manner, the first reaction zone canoperate at higher temperature and pressure conditions that favor ringopening but performs only a portion of the normal to isoparaffinconversion. The two stage heating of the combined feed, e.g., asprovided by exchangers 26 and 24, facilitates the use of highertemperatures therein in a first isomerization reactor. Once cyclichydrocarbon rings have been opened by initial contact with the catalyst,the final reactor stage may operate at temperature conditions that aremore favorable for isoalkane equilibrium.

Another benefit of using two reactors is that it allows partialreplacement of the catalyst system without taking the isomerization unitoff stream. For short periods of time, during which the replacement ofcatalyst may be necessary, the entire flow of reactants may be processedthrough only one reaction vessel while catalyst is replaced in theother.

Whether operated with one or two reaction zones, the effluent of theprocess will enter separation facilities for the recovery of anisoalkane product. At minimum, the separation facilities divide thereaction zone effluent into a product stream comprising C₅ and heavierhydrocarbons and a gas stream which is made up of C₃ lighterhydrocarbons and hydrogen. To the extent that C₄ hydrocarbons arepresent, the acceptability of these hydrocarbons in the product streamwill depend on the blending characteristics of the desired product, inparticular vapor pressure considerations. Consequently, C₄ hydrocarbonsmay be recovered with the heavier isomerization products or withdrawn aspart of the overhead or in an independent product stream. Suitabledesigns for rectification columns and separator vessels to separate theisomerization zone effluent are well known to those skilled in the art.

When hydrogen is received for recycle from the isomerization zoneeffluent, the separation facilities, in simplified form, can consist ofa product separator and a stabilizer. The product separator operates asa simple flash separator that produces a vapor stream rich in hydrogenwith the remainder of its volume principally comprising C₁ and C₂hydrocarbons. The vapor stream serves primarily as a source of recyclehydrogen which is usually returned directly to the hydrogenationprocess. The separator may contain packing or other liquid vaporseparation devices to limit the carryover of hydrocarbons. The presenceof C₁ and C₂ hydrocarbons in the vapor stream do not interfere with theisomerization process, therefore, some additional mass flow for thesecomponents is accepted in exchange for a simplified column design. Theremainder of the isomerization effluent leaves the separator as a liquidwhich is passed on to a stabilizer, typically a trayed column containingapproximately 30 trays. The column will ordinarily contain condensingand reboiler loops for the withdrawal of a light gas stream comprisingat least a majority of the remaining C₃ hydrocarbons from the feedstream and a light bottoms stream comprising C₅ and heavierhydrocarbons. Normally when the isomerization zone contains only a smallquantity of C₄ hydrocarbons, the C₄'s are withdrawn with the light gasstream. After caustic treatment for the removal of chloride compounds,the light gas stream will ordinarily serve as fuel gas. The stabilizeroverhead liquid, which represents the remainder of the isomerizationzone effluent passes back to the fractionation zone as recycle input.

A simplified flow scheme for use without hydrogen recycle stream wasdescribed in the FIGURE. In the arrangement of the FIGURE, all of theexcess hydrogen from the isomerization zone is taken with the overheadstream from the stabilizer drum or receiver. Since, as a preconditionfor use of this arrangement, the amount of hydrogen entering thestabilizer is low, the rejection of hydrogen with the fuel gas streamdoes not significantly increase the loss of product hydrocarbons.

In order to more fully illustrate the process, the following theoreticalexample is presented to demonstrate the operation of the processutilizing the flow scheme of the FIGURE. All of the numbers identifyingvessels and lines correspond to those given in the FIGURE. The Tableprovides illustrative compositions of streams of the process. The Tableis merely an example, and stream compositions may vary from those shown.

A C₅ plus naphtha fresh feed having a composition shown in the Tableenters through line 10 and is heat exchanged with the hydrogenation zoneeffluent before being passed through sulfur guard bed 12 to removesulfur components. The sulfur-free feed is conducted in line 13 to lowpressure drier 11 to remove water. Feed in line 13 may be combined withrecycle normal alkanes in line 64 from deisohexanizer 58 prior to beingdried in low pressure drier 11. Furthermore, if some or all of the feedis light reformate, it is expected that the light reformate will alreadybe sulfur-free and sulfur guard bed 12 may be bypassed or eliminated.Optional line 70, shown as a dashed line, shows light reformate feedbypassing sulfur guard bed 12. Reducing the amount of material passingthrough the sulfur guard bed may result in a smaller guard bed beingrequired thus reducing costs. Hydrogen in line 14 is dried in drier 16and combined with dried feed in line 15 to form a combined feed.

Combined feed 18 is passed through a series of heat exchangers such asexchangers 24, 25 and 26 to heat the feed to a temperature of 149° C. to204° C. (300° to 400° F.) which then enters the hydrogenation reactor ata pressure of 3450 kPa(g) (500 psig). In the hydrogenation reactor, thecombined feed is contacted with a catalyst comprising a platinum metalon a chlorided platinum alumina support at an LHSV of 20. Contact of thecombined feed with the hydrogenation catalyst produces a saturated feedthat is withdrawn by line 34 and has no more than about 0.5 wt.-%benzene. The hydrogenation zone heats the saturated feed to atemperature of 177 to 274° C. (3500 to 525° F.). Since the temperaturerequired for the isomerization zone is less than the temperature of thesaturated feed, the saturated feed is heat exchanged with the combinedfeed in line 18 and with the fresh feed in line 10. If necessary, thesaturated feed may also be cooled. The saturated feed in line 34 ispassed on to the isomerization zone at a pressure of 3240 kPa(g) (470psig). Perchloroethylene is added to the saturated feedstream at a rateof 150 wt. ppm which then enters the reactor train 30 and 22 of theisomerization zone. In the isomerization zone, the saturated feed streamcontacts an alumina catalyst such as one having 0.25 wt.-% platinum and5.5 wt.-% chloride. The converted isomerization zone feed passed out ofthe reactor train in line 20 at a temperature of 93 to 204° C. (200 to400° F.) and a pressure of 3100 kPa(g) (450 psig) and has an exemplarycomposition as shown in the Table.

After heat exchange with combined feed 18, cooled isomerization zoneeffluent in line 36 enters the stabilizer column 38 for the recovery ofthe product and removal of light gases. Column 38 has, for example, 30trays and the feed may enter above tray 15. Column 38 splits theisomerization zone effluent into an overhead 42 which is cooled andcondensed 44 to provide a recycle 46 and a fuel gas stream 50. Becauseof the chloride in the stream, the fuel gas stream 50 is passed throughscrubber 52 to remove any chloride and provide a scrubbed fuel gasstream 56. Spent caustic is removed from scrubber 52 in stream 54. Anisomerization zone product 40 is withdrawn from the bottom of stabilizercolumn 38 and has the exemplary composition shown in the Table.Isomerization zone product 40 is passed to deisohexanizer 58 to separatelow octane normal and monomethyl alkanes into stream 64 which may berecycled to combine with the feed stock in line 13. The pentanes,dimethyl-butanes, and some monomethyl alkanes removed in DIH overhead 60are combined with the C6 naphthenes and C7+ in DIH bottoms 62 to formthe process product stream 66.

This example demonstrates the ability of the process to saturate benzeneusing a flow scheme that allows low pressure feedstock driers andrequires no condensing of the feed that would also require a receiverwith hydrogen venting and additional pumps. The combined feed is heatexchanged with the effluents of the isomerization reactors and thebenzene saturation reactor, and the benzene saturation effluent is alsoheat exchanged with the fresh feed. All values in the table are merelyexemplary of one embodiment, and the compositions of the stream may varywith different applications.

TABLE Stream Compositions in kmol/hr Stream Number 10 14 34 20 56 40Hydrogen 0.0 46.0 14.4 5.0 5.0 0.0 C1-C4 0.1 0.1 0.2 3.7 3.6 0.1Isopentane 10.4 0.0 10.4 20.8 0.2 20.6 Normal Pentane 15.6 0.0 15.6 7.10.0 7.1 Cyclopentane 1.7 0.0 1.7 1.4 0.0 1.4 Dimethylbutanes 4.5 0 4.520.0 0 20.0 Methylpentanes 21.4 0 21.4 24.1 0 24.1 Normal Hexane 19.90.0 19.9 6.2 0.0 6.2 Methylcyclopentane 7.9 0.0 7.9 7.1 0.0 7.1Cyclohexane 5.9 0.0 16.4 8.3 0.0 8.3 Benzene 10.7 0.0 0.2 0.0 0.0 0.0C7+ 1.9 0 1.9 3.2 0 3.2 Total 100 46.1 114.5 106.9 8.8 98.1

1. A process for the hydrogenation and decyclization of benzene and the isomerization of C₅-C₆ paraffins with a feedstock comprising C₅-C₆ paraffins and at least 1 wt.-% benzene, said process comprising: (a) passing the feedstock, without venting hydrogen or condensing, to a drier, to remove water and thereby dry the feedstock and generate a dried feedstock comprising less than 0.5 wt.-% water; (b) combining the dried feedstock with a hydrogen-rich gas stream to produce a combined feed; (c) passing the combined feed, at a temperature of from about 38 to about 232° C. (about 100 to about 450° F.), to a hydrogenation zone and contacting said combined feed with a hydrogenation catalyst at hydrogenation conditions to saturate benzene and generate a hydrogenation zone effluent having a temperature in the range of about 149 to about 288° C. (300 to 550° F.) and comprising less than 1.5 wt. % benzene; (d) adjusting the temperature of the hydrogenation zone effluent to a range of about 104 to about 204° C. (220 to 400° F.) by at least heat exchanging the hydrogenation zone effluent with the combined feed; (e) passing at least a portion of the hydrogenation zone effluent to an isomerization zone and contacting with an isomerization catalyst at isomerization and decyclization conditions to isomerize C₅-C₆ paraffins and decyclize benzene; (f) recovering an isomerate product from the isomerization zone; (g) separating the isomerate product to remove C₄ and lighter hydrocarbons and generate a stabilizer bottoms stream; and (h) recovering the stabilizer bottoms stream as product.
 2. The process of claim 1 further comprising passing at least a portion of the feedstock through a sulfur guard bed to remove sulfur-containing components prior to passing the feedstock to the drier.
 3. The process of claim 2 further comprising combining a second feed with the feedstock after passing the feedstock through the sulfur guard bed and prior to passing to the drier.
 4. The process of claim 1 wherein the feedstock comprises about 1 to about 25 wt.-% benzene.
 5. The process of claim 1 wherein the combined feed to the hydrogenation zone is heated to about 127 to 232° C. (260 to 450° F.).
 6. The process of claim 1 wherein the hydrogenation zone effluent has from about 0.01 to about 1.5 wt.-% benzene.
 7. The process of claim 1 wherein the hydrogenation zone effluent has a temperature in the range of about 177 to 274° C. (350 to 525° F.).
 8. The process of claim 1 wherein the hydrogenation zone effluent is adjusted to a temperature in the range of about 38 to 260° C. (100 to 500° F.).
 9. The process of claim 1 wherein the hydrogen-rich gas stream is mixed with said feedstock to produce a hydrogen to hydrocarbon ratio in the range of 0.1 to 2 in the combined feed.
 10. The process of claim 1 wherein said feedstock additionally comprises C₄ paraffins.
 11. The process of claim 1 wherein the hydrogenation conditions include a pressure of from 1380 to 4830 kPa(g) (200 to 700 psig), a liquid hourly space velocity of from 1 to 40 and a hydrogen to hydrocarbon ratio of from 0.1 to
 2. 12. The process of claim 1 wherein said isomerization catalyst comprises a chlorided platinum catalyst on alumina support.
 13. The process of claim 1 wherein a chloride concentration of from 30-300 ppm is maintained in the isomerization zone by injecting a chloride compound into the hydrogenation zone effluent.
 14. The process of claim 1 wherein the isomerization zone includes at least two reactors in series, the first reactor is operated at conditions to open saturated hydrocarbon rings, said conditions including a temperature in excess of 143° C. (290° F.) and a pressure of at least 1380 to 4690 kPa(g) (200 to 680 psig) and the second reactor in the series is operated at conditions to increase the concentration of C₅-C₆ isoalkanes including a temperature in the range of 38 to 177° C. (104 to 350° F.).
 15. The process of claim 1 wherein the drier of step 1(a) is a low pressure drier.
 16. The process of claim 1 wherein the hydrogenation catalyst comprises a platinum group metal component on a solid support. 